Slurry phase polymerisation process

ABSTRACT

Polymerisation process in which polyethylene is produced in slurry in a polymerisation reactor in the presence of a Ziegler Natta catalyst and an activator, and a stream or slurry containing the polymer is withdrawn from the reactor and transferred to a flash tank operating at a pressure and temperature such that at least 50 mol % of the liquid or non-polymer component of the stream entering the flash tank or slurry is withdrawn from the flash tank as a vapour and at least 98 mol % of the vapour withdrawn from the flash tank is capable of being condensed at a temperature of between 15 and 50° C., without compression. A by-product suppressor, which reduces the amount of by-product formed per unit of polyethylene produced by at least 10%, compared with an identical polymerisation process where the by-product suppressor is not present, is used in the reactor. The molar ratio of the by-product suppressor added to the reactor to titanium added to the reactor is between 0.2 and 1.

This application is the U.S. national phase of International ApplicationNo. PCT/EP2008/061369 filed 29 Aug. 2008 which designated the U.S. andclaims priority to European Application No. 07253479.5 filed 3 Sep.2007, the entire contents of each of which are hereby incorporated byreference.

The present invention is concerned with olefin polymerisation in slurryphase reactors.

Slurry phase polymerisation of olefins is well known wherein an olefinmonomer and optionally olefin comonomer are polymerised in the presenceof a catalyst in a diluent in which the solid polymer product issuspended and transported.

Polymerisation is typically carried out at temperatures in the range50-125° C. and at pressures in the range 1-100 bara. The catalyst usedcan be any catalyst typically used for olefin polymerisation such aschromium oxide, Ziegler-Natta or metallocene-type catalysts.

Polymerisation is typically carried out in loop reactors, which are of acontinuous tubular construction comprising at least two, for examplefour, vertical sections and at least two, for example four horizontalsections. The heat of polymerisation is typically removed using indirectexchange with a cooling medium, preferably water, in jackets surroundingat least part of the tubular loop reactor. The volume of each loopreactor of a multiple reactor system can vary but is typically in therange 10-200 m³, more typically 50-120 m³. The loop reactors employed inthe present invention are of this generic type.

Typically, in the slurry polymerisation process of polyethylene forexample, the slurry in the reactor will comprise the particulatepolymer, the hydrocarbon diluent(s), (co) monomer(s), catalyst, chainterminators such as hydrogen and other reactor additives In particularthe slurry will comprise 20-80 weight percent (based on the total weightof the slurry) of particulate polymer and 80-20 weight percent (based onthe total weight of the slurry) of suspending medium, where thesuspending medium is the sum of all the fluid components in the reactorand will comprise the diluent, olefin monomer and any additives; thediluent can be an inert diluent or it can be a reactive diluent inparticular a liquid olefin monomer; where the principal diluent is aninert diluent the olefin monomer will typically comprise 2-20,preferably 4-10 weight percent of the slurry.

The slurry is pumped around the relatively smooth path endless loopreaction system at fluid velocities sufficient to maintain the polymerin suspension in the slurry and to maintain acceptable cross-sectionalconcentration and solids loading gradients. Slurry is withdrawn from thepolymerisation reactor containing the polymer together with the reagentsand inert hydrocarbons, all of which mainly comprise inert diluent andunreacted monomer. The product slurry comprising polymer and diluent,and in most cases catalyst, olefin monomer and comonomer can bedischarged intermittently or continuously, optionally usingconcentrating devices such as hydrocyclones or settling legs to minimisethe quantity of fluids withdrawn with the polymer. The reagents andhydrocarbons need to be separated from the polymer and recovered foreconomic, safety and environmental reasons, and many processes forachieving this are known in the art. These processes generally involvedepressurising and devolatilising the polymer-containing stream after ithas been withdrawn from the polymerisation reactor. The diluent andunreacted monomer can then be recompressed back into liquid form andrecycled back into the reactor, whilst the solid polymer can betransferred for further processing.

A well-known disadvantage of the above process for separating andrecycling the liquid components withdrawn from the reactor with thepolymer is that recompressing them back into liquid form after they havebeen vaporised in the separation process requires considerable energy.Thus for single reactor operations, alternatives have been proposed suchas in WO 99/47251 where the majority of the liquid components withdrawnwith the polymer are separated in a flash tank at a temperature andpressure such that they can be recondensed just by cooling, withoutrecompression. The remaining liquid components not removed by thisprocess are separated in a second flash tank operating at a lowerpressure, and these need to be recompressed in order to be recycled. Theadvantage of this process, which is referred to hereinafter as a “mediumpressure flash” process, is that only a small proportion of thevaporised liquid components need to be recompressed in order to berecondensed.

Whilst the above “medium pressure flash” process has been found to besuitable in single reactor polymerisations, it is nevertheless dependentfor its success on the composition of the slurry withdrawn from thereactor system and depressurised. If the content of unreacted monomer,hydrogen and other light components in the slurry withdrawn from thereactor is too high, it will not be possible to flash at a temperatureand pressure which would allow both vaporisation of a substantialportion of those components and also recondensation thereof (using aneconomically viable coolant such as water) without compression. In sucha case it would still be necessary to recompress the vaporisedcomponents in order to recycle them to the reactor.

We have found that it is possible to improve control the composition ofthe slurry withdrawn from the reactor so as to ensure that the above“medium pressure flash” process can operate without the need forcompression, by adding a by-product suppressor to the reactor.

Accordingly, in a first aspect the invention provides a polymerisationprocess in which polyethylene is produced in slurry in a polymerisationreactor in the presence of a Ziegler Natta catalyst and an activator,and a stream or slurry containing the polymer is withdrawn from thereactor and transferred to a flash tank operating at a pressure andtemperature such that at least 50 mol % of the liquid or non-polymercomponent of the stream entering the flash tank or slurry is withdrawnfrom the flash tank as a vapour and at least 98 mol %, more preferablyat least 98.5 mol %, and most preferably at least 99.5 mol %, of thevapour withdrawn from the flash tank is capable of being condensed at atemperature of between 15 and 40° C. without compression, wherein aby-product suppressor, which reduces the amount of by-product formed perunit of polyethylene produced by at least 10%, and preferably theabsolute amount of by-product formed in the reactor by at least 10%,compared with an identical polymerisation process where the by-productsuppressor is not present, is used in the reactor.

Usually the amount of by-product formed is measured as a production ratein kg/h, and hence the reduction is usually calculated by reference tosuch a production rate.

A “by-product suppressor” is a compound which suppresses the formationof by-products during the polymerisation reaction, by which is meantthat it reduces the amount of by-product formed per unit of polyethyleneproduced, such that the ratio of by-product to polyethylene is reducedby at least 10%, preferably at least 20% (compared with an identicalpolymerisation where the by-product suppressor is not present).Preferably the by-product suppressor also reduces the absolute amount ofby-product produced by at least 10%, preferably at least 20% (comparedwith an identical polymerisation where the by-product suppressor is notpresent). We have found that by introducing such a compound into thereactor, it is possible to reduce the proportion of components having amolecular weight below 50 in the reactor and hence also in the flashtank. As explained above, an excess of such light components in theslurry which is passed to the flash tank can make it difficult torecondense the flashed vapour without first compressing it. Usually theby-product suppressor acts by reducing the rate of formation in thereactor of one or more products other than polyethylene by at least 10%,preferably at least 20%, (compared with an identical polymerisationwhere the by-product suppressor is not present).

Preferably the by-product suppressor also enhances the polymerisationactivity. In particular, it is preferred that the by-product suppressorincreases the polymerisation activity by at least 10%, preferably atleast 20%, compared with an identical polymerisation where theby-product suppressor is not present.

Preferably at least 80 mol %, more preferably 90 mol %, most preferably95 mol % of the liquid component of the slurry is withdrawn from theflash tank as a vapour.

Components of the slurry which have a molecular weight below 50 g/mol(“C_(lights)”) are typically all hydrocarbon components containing 3carbon atoms or less, including ethylene, propylene, methane, ethane,propane and hydrogen. It is preferred that the by-product suppressorreduces the concentration at least one of the components in the slurrywhich has a molecular weight below 50 by at least 5 mol %, preferably atleast 10 mol %.

In order to ensure that at least 98 mol % of the vapour withdrawn fromthe flash tank is capable of being condensed without compression, it ispreferred that the concentration in the slurry entering the flash tankof components having a molecular weight below 50 g/mol, C_(lights) (mol%), satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) where T_(c)and P_(c) are respectively the temperature (in ° C.) and pressure (MPag) at the location where the vapour withdrawn from the flash tank iscondensed, and C_(H2) and C_(Et) are the molar concentrations in theflash tank of hydrogen and ethylene respectively.

It is preferred that the concentration of components having a molecularweight below 50 g/mol in the slurry entering the flash tank iscontrolled solely by controlling that concentration in the reactor.Accordingly it is preferred that the concentration in the reactor ofcomponents having a molecular weight below 50 g/mol also satisfies theequation C_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et))where C_(lights), C_(H2), and C_(Et) in this case are the concentrationsof components having a molecular weight below 50 g/mol, hydrogen andethylene respectively in the reactor. More preferably the concentrationof components having a molecular weight below 50 g/mol in the reactor isthe same as the concentration of components having a molecular weightbelow 50 g/mol entering the flash tank.

The by-product suppressor is preferably something which reduces theconcentration and/or suppresses the formation of one or more compoundshaving a molecular weight below 50 g/mol. An example of a by-productsuppressor is a halogenated hydrocarbon, and more particularly achloromethane of the formula CH_(x)Cl_(4-x) where x is an integer from 1to 3.

In an alternative aspect, the present invention provides apolymerisation process in which polyethylene is produced in slurry in apolymerisation reactor in the presence of a Ziegler Natta catalyst andan activator, and a stream or slurry containing the polymer is withdrawnfrom the reactor and transferred to a flash tank operating at a pressureand temperature such that at least 50 mol % of the liquid or non-polymercomponent of the stream entering the flash tank or slurry is withdrawnfrom the flash tank as a vapour and at least 98 mol %, more preferablyat least 98.5 mol %, and most preferably at least 99.5 mol %, of thevapour withdrawn from the flash tank is capable of being condensed at atemperature of between 15 and 40° C. without compression, wherein ahalogenated hydrocarbon, preferably a chloromethane of the formulaCH_(X)Cl_(4-x) where x is an integer from 1 to 3, is used in thereactor.

In both aspects of the invention, it is preferred that a streamcontaining the polymer is withdrawn from the reactor and transferred toa flash tank operating at a pressure and temperature such that at least50 mol % of the non-polymer component of the stream entering the flashtank is withdrawn from the flash tank as a vapour. Usually the streamexits the reactor as a slurry of which the non-polymer component in inthe liquid phase. By the time the stream enters the flash tank however,most if not all of the non-polymer component is in the vapour phase.

In both aspects of the invention, the most preferred chloromethane ischloroform, CHCl₃. The use of a halogenated hydrocarbon is preferredbecause it has the dual effect of both suppressing ethane formation andalso enhancing polymerisation activity. Ethane formation adds to theconcentration of light reagents in the reactor, and therefore if notremoved would make it more difficult to avoid the need to compress thevapour flashed from the flash tank in order to recondense it. In factethane is normally purged from the reactor, but this results in anassociated loss of ethylene. Reducing the amount of ethane formed in thefirst place permits a reduction in the amount of ethane purged, andhence a reduction in the amount of ethylene lost. Ethane formation canbe particularly significant when making low molecular weight polymers,particularly if hydrogen is present. It can also be desirable to boostthe activity of the catalyst since the high hydrogen concentration cancontribute to a reduction in polymerisation activity. Halogenatedhydrocarbons such as chloroform can therefore provide a double benefit,by boosting activity boost and also minimising the concentration ofC_(lights) in the second reactor.

The amount of by-product suppressor added is usually based on the amountof Ziegler-Natta catalyst, and is preferably such that the molar ratioof the by-product suppressor added to the reactor to titanium added tothe reactor is greater than 0.1, preferably between 0.2 and 1.

In one embodiment of the invention, the concentration of componentshaving a molecular weight below 50 g/mol in the slurry entering theflash tank is additionally controlled by a post-polymerisationtreatment, in which the conditions between the reactor and the flashtank are controlled to ensure that further polymerisation takes placebefore entering the flash tank, thereby consuming some of the componentshaving a molecular weight below 50 g/mol. The object of thepost-polymerisation stage is to reduce the concentration of componentshaving a molecular weight below 50 g/mol in the slurry whilst minimizingthe effect on the final product properties. The residence time,temperature and slurry velocity are controlled to achieve the reductionrequired in the concentration of these components whilst avoidingblockage or fouling in the post-polymerisation zone. In order to achievethe required degree of polymerisation, it is preferred that in thisembodiment the residence time of the slurry in the post-polymerisationzone between the reactor and the flash tank is at least 20 seconds,preferably at least 5 minutes and more preferably between 10 and 30minutes. Typically residence times of 15-25 minutes are employed. Thepost-polymerisation zone may be provided in the form of an agitatedtank; however it is most preferred that it is in the form of an expandeddiameter transfer line between the second reactor and the flash tank,which provide essentially plug flow rather than “continuous stirredtank-like” reaction conditions. The volume of the expanded diameterportion of the transfer line is typically between 1 m³ and 35 m³,preferably between 5 and 25 m³. Its preferred length to internaldiameter is preferably between 100 and 1000, more preferably between 250and 600. It is also preferred that the internal diameter is between 200and 1000 mm, most preferably between 500 and 750 mm. The residence timeis preferably installed upstream of a hydrocyclone, the slurry pressurelet-down valve and the in-line slurry heater.

In an alternative embodiment of the invention, the concentration ofcomponents having a molecular weight below 50 g/mol in the slurryentering the flash tank is additionally controlled either by introducingadditional liquid into the slurry as it passes through the transfer linebetween the reactor and the flash tank, and/or by adjusting the solidsconcentration of the slurry as it passes through said transfer line. Theadditional liquid which may be introduced is usually inert diluenthaving a lower concentration of components with a molecular weight below50 (C_(lights)) than that of the slurry withdrawn from the reactor, andpreferably a C_(lights) level that is 25-50% of the C_(lights) level inthe slurry itself. The C_(lights) of the additional liquid is preferablyless than 1 mol %, most preferably 0 mol %. The solids concentration ofthe slurry may be adjusted by passing the slurry through a hydrocyclonelocated in the transfer line: the solids-rich stream is passed to theflash tank and the solids-lean stream is recycled either upstream in thetransfer line or back to the reactor.

The above embodiments of the invention involving post-reactor treatmentof the slurry leaving the reactor make it possible to increase theactivity in the reactor by increasing the concentration therein ofreactants such as ethylene as well as by adding a by-product suppressorwhich may additionally enhance activity, without compromising theability to avoid recompression of the liquid vaporised in the firstflash tank.

In all embodiments of the invention, and referring either to the flashtank and/or the reactor as discussed above, it is generally preferredthat the concentration of components having a molecular weight below 50satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) whereC_(lights), C_(H2), C_(Et), P and T_(c) are as defined previously andrefer either to the reactor or the flash tank depending on theparticular embodiment of the invention.

It is preferred that the polyethylene made in the reactor is amultimodal polyethylene having a shear ratio of at least 15, generallybetween 15 and 50, and preferably between 21 and 35. By “shear ratio” isthe ratio of the high load melt index HLMI of the polyethylene to theMI₅ of the polyethylene. The HLMI and MI₅ are measured according to ISOStandard 1133 at a temperature of 190° C. using loads of 21.6 kg and 5kg respectively. MI₂ is similarly measured but using a load of 2.16 kg.

The HLMI of the multimodal polyethylene exiting the reactor ispreferably between 1 and 100 g/10 min, and more preferably between 1 and40 g/10 min.

Typically the multimodal polyethylene is made in at least two reactorsin series, one of which is the reactor to which the by-productsuppressor is added. Although it may be made in more than two reactors,it is most preferred that the multimodal polyethylene is a bimodalpolymer made in two reactors in series, with the by-product suppressorbeing added to either reactor. However this does not exclude thepossibility that up to 10 wt % of a third polymer may be made betweenthe two reactors. It also does not exclude the possibility ofpolymerisation taking place prior to the first reactor, for example in aprepolymerisation reaction. The first polymer made in the first reactormay be a low molecular weight (LMW) polymer and the second polymer madein the second reactor may be a high molecular weight (BMW) polymer, orvice versa. In one embodiment, 30-70 wt % and more preferably 40-60 wt %of a low molecular weight (LMW) polymer is made in the first reactor,and 70-30 wt % and more preferably 60-40 wt % of a high molecular weight(HMW) polymer is made in the second reactor. The most preferred range ofratios of the HMW and LMW polymers is 45-55 wt % to 55-45 wt %. Theterms “low molecular weight” (LMW) and “high molecular weight” (BMW) areintended to be relative terms, in that the LMW polymer has a lowermolecular weight than the HMW polymer; there is no limit on the absolutemolecular weights which may be made in each reactor.

In a preferred embodiment, the HMW polymer is made in suspension in thefirst reactor and the LMW polymer is made in suspension in the secondreactor in the presence of the first polymer, the ratios of each polymerpreferably being 30-70 wt % and 70-30 wt % respectively. The followingrequirements apply to this embodiment only.

In this embodiment of the invention (“HMW-LMW”), it is preferred thatthe by-product suppressor is added to the second LMW reactor. It is alsopreferably ensured that the concentration of components having amolecular weight below 50 in the second LMW reactor satisfies theequation C_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) byensuring the ratio of the average activity in the second LMW reactor tothe average activity in the first HMW reactor is from 0.25 and 1.5.Average activity is typically higher in the first reactor (where acopolymer is usually made to obtain the HMW product) than in the secondreactor (where a homopolymer or high density polymer is usually made toobtain the LMW product), and we have found that as a consequence theratio of average activities between the reactors has to be controlledwithin these ranges in order to control the concentration of lightcomponents in the second reactor. The average activity in each reactoris defined as the rate of polyethylene produced in the reactor(kgPE/hr)/[ethylene concentration in the reactor (mol %)×residence timein the reactor (hours)×feed rate of catalyst into the reactor (g/hr)].If no additional catalyst is added to the second reactor, whencalculating the ratio of average activities the flow rate of catalyst inthe two reactors is considered to be the same. If additional catalyst isadded to the second reactor, the flow rate into the second reactor isconsidered to be the sum of the flowrate of catalyst from the firstreactor plus the flowrate of additional fresh catalyst added directlyinto the second reactor. Alternatively, activity in each reactor may becalculated based on catalyst residues in the polymer produced in eachreactor, as is well known, and the activity ratio calculated from this.

The residence time is defined as the mass of the polymer in the reactor(kg)/the output rate of polymer from the reactor (kg/hr). In a casewhere polymer is recycled back into the reactor, for example when ahydrocyclone is employed downstream of the reactor, the output rate ofpolymer is the net output rate (ie polymer withdrawn less polymerrecycled).

Preferably the overall productivity of the process is at least 10 kgpolyethylene/g catalyst, preferably greater than 15 kg polyethylene/gcatalyst, most preferably greater than 20 kg polyethylene/g catalyst.

In order to achieve the above ratio of average activities, it ispreferred that the ratio of ethylene concentration (in mol %) in thesecond reactor to that in the first reactor is 5 or less. Preferably theratio of ethylene concentration in the second reactor to that in thefirst reactor is 3 or less, and more preferably 2.5 or less. Mostpreferably both ethylene concentration ratio and average activity ratiorequirements are satisfied together.

In this HMW-LMW embodiment it is preferred that the actual concentrationof ethylene in the second reactor is less than 8 mol %. However in orderto ensure a satisfactory level of productivity, it is also preferredthat the ethylene concentration is greater than 1.5 mol %, preferablygreater than 2 mol %. The concentration of hydrogen in the secondreactor is preferably less than 5 mol %, more preferably less than 3 mol%. The ratio of hydrogen to ethylene is preferably 0-0.5 mol/mol.

Usually each of the reactors has an internal volume greater than 10m³,more commonly greater than 25 m³ and in particular greater than 50 m³.Typical ranges are 75-200 m³ and more particularly 100-175 m³. In oneversion of the HMW-LMW embodiment of the invention the volumes of eachof the reactors employed differ by less than 10%, and it is preferredthat all of the volume, length and diameter of the reactors employedeach independently differ by less than 10%. Most preferably in thisversion of the embodiment, the reactors have the same dimensions.

Thus in the HMW-LMW embodiment of the invention, in the case where thereactors differ in volume by no more than 10 vol %, it is preferred tobalance the activities between the reactors and the respective coolingcapacities by maintaining the temperature of the first reactor between60 and 80° C., preferably no higher than 75° C. It is also preferredthat the ratio of solids concentration in the first reactor to that inthe second reactor is maintained at less than 1.0, preferably between0.6 and 0.8, as this also assists in maintaining the balance of averageactivity between the two reactors within the desired range.

Generally in the HMW-LMW embodiment of the invention, the solidsconcentration in the final LMW reactor is at least 35 wt %, mostpreferably between 45 wt % and 60 wt % and the solids concentration inthe HMW reactor is between 20 wt % and 50 wt %, more preferably between25 wt % and 35 wt %. The solids concentration is the weight of polymerrelative to the total weight of the slurry. In this case it is preferredto concentrate the solids transferred from the first reactor to thesecond reactor using a settling zone and/or hydrocyclone. Acomonomer-free diluent stream may be introduced upstream of thehydrocyclone to reduce the proportion of comonomer transferred to thedownstream reactor, thus increasing the density of the polymer producedin the LMW reactor.

By maintaining the preferred ratio of average activity and ethyleneconcentration ratio between the two reactors in the HMW-LMW embodiment,it is possible to achieve high overall space time yields (defined asproduction of polymer in kg/h per unit volume of reactor) and activitieswhilst still observing the C_(lights) requirements of the invention inthe flash tank. The average space time yield in all reactors combinedmay be maintained at greater than 100 kg/m³/h, more preferably greaterthan 150 kg/m³/h, and most preferably greater than 200 kg/m³/h.

However, where a plant has been designed to operate a single catalyst orproduct type, the volume and dimensions of each reactor may be optimisedindividually for the principal grades to be produced, and thus the tworeactors may be of different volumes and dimensions. These differentdimensions can be utilised in order to obtain the desired balance ofaverage activity between the two reactors in accordance with theinvention, thus providing greater freedom to vary other reactionparameters. Thus in order to achieve the desired ratio of averageactivity between the first (HMW) and second (LMW) reactors, in analternative embodiment of the HMW-LMW aspect of the invention, the first(HMW) reactor may have a space time yield (defined as production ofpolymer in kg/h per unit volume of reactor) greater than 150 kg/m³/h,more preferably greater than 200 kg/m³/h, and most preferably greaterthan 250 kg/m³/h. It is also preferred in this case that the ratio ofspace time yield in the first (BMW) reactor to the second (LMW) reactoris greater than 1.2, most preferably greater than 1.5. This may beachieved by designing the first (HMW) reactor with a volume that is nomore than 90%, preferably between 30-70%, and more preferablyapproximately 40-60%, of the volume of the second (LMW) reactor. In thiscase it is preferred that the ratio of length to diameter (LID) of thefirst reactor is greater than 500, preferably between 750 and 3000, andmost preferably greater than 800;additionally or alternatively, theratio of L/D of the first reactor to LID of the second reactor isgreater than 1.5, most preferably greater than 2. In this reactorconfiguration, the average space time yield in all reactors combined maybe maintained at greater than 150 kg/m³/h, more preferably greater than200 kg/m³/h, and most preferably greater than 300 kg/m³/h.

Alternatively, the LMW polymer is made in the first reactor and the HMWpolymer in the second reactor. In this case, it is generally preferredthat the by-product suppressor is added to the first LMW reactor. Thefollowing requirements apply to this embodiment only.

In this embodiment of the invention (“LMW-HMW”), it is preferablyensured that the concentration of components having a molecular weightbelow 50 in the second reactor satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) by ensuringthe ratio of average activity in the second HMW reactor to averageactivity in the first LMW reactor is from 1.5 to 0.25. The averageactivity in each reactor is defined as the rate of polyethylene producedin the reactor (kgPE/hr)/[ethylene concentration in the reactor (mol%)×residence time in the reactor (hours)×feed rate of catalyst into thereactor (g/hr)]. If no additional catalyst is added to the secondreactor, when calculating the ratio of average activities the flow rateof catalyst in the two reactors is considered to be the same. Ifadditional catalyst is added to the second reactor, the flow rate intothe second reactor is considered to be the sum of the flowrate ofcatalyst from the first reactor plus the flowrate of additional freshcatalyst added directly into the second reactor. Alternatively, activityin each reactor may be calculated based on catalyst residues in thepolymer produced in each reactor, as is well known, and the activityratio calculated from this.

The residence time is defined as the mass of the polymer in the reactor(kg)/the output rate of polymer from the reactor (kg/hr). In a casewhere polymer is recycled back into the reactor, for example when ahydrocyclone is employed downstream of the reactor, the output rate ofpolymer is the net output rate (ie polymer withdrawn less polymerrecycled).

This embodiment of the present invention is particularly applicable whenthe polymerisation catalyst is a Ziegler-Natta catalyst, especially ifthe overall productivity of the process is at least 10 kg polyethylene/gcatalyst, preferably greater than 15 kg polyethylene/g catalyst, mostpreferably greater than 20 kg polyethylene/g catalyst.

In order to achieve the above ratio of average activity, it is preferredthat the ratio of ethylene concentration (in mol %) in the secondreactor to that in the first reactor is 5 or less. Most preferably bothethylene concentration ratio and ratio of average activity requirementsare satisfied together. Preferably the ratio of ethylene concentrationin the second reactor to that in the first reactor is 3 or less, andmore preferably 2 or less.

In this LMW-HMW embodiment it is preferred that the actual concentrationof ethylene in the second reactor is less than 8 mol %. However in orderto ensure a satisfactory level of productivity, it is also preferredthat it is greater than 2 mol %. The ethylene concentration ispreferably between 2 and 5 mol %. The concentration of hydrogen in thesecond reactor is preferably less than 5 mol %, more preferably lessthan 3 mol %.

Usually each of the reactors has an internal volume greater than 10 m³,more commonly greater than 25 m³ and in particular greater than 50 m³.Typical ranges are 75-200 m³ and more particularly 100-175 m³. In oneversion of the LMW-HMW embodiment of the invention the volume of thereactors employed each independently differ by less than 10%, and it ispreferred that all of the volume, length and diameter of the reactorsemployed each independently differ by less than 10%. Most preferably inthis version of the embodiment, the reactors have the same dimensions.Thus in the LMW-HMW embodiment of the invention, in the case where thereactors differ in volume by no more than 10 vol %, it is preferred tobalance the activities between the reactors and the respective coolingcapacities by maintaining the temperature of the first reactor between70 and 110° C., preferably between 80 and 100° C. It is also preferredthat the ratio of solids concentration in the first reactor to that inthe second reactor is maintained at between 0.8 and 1.2, preferablybetween 0.9 and 1.0, as this also assists in maintaining the balance ofaverage activity between the two reactors within the desired range.

Generally in the LMW-HMW embodiment of the invention, the solidsconcentration in each reactor is at least 35 wt %, most preferablybetween 45 wt % and 55 wt %. The solids concentration is the weight ofpolymer relative to the total weight of the slurry. In this case it ispreferred to concentrate the solids transferred from the first reactorto the second reactor using a settling zone and/or hydrocyclone. Ahydrogen-free diluent stream may be introduced upstream of thehydrocyclone to reduce the proportion of hydrogen transferred to thedownstream reactor. It is most preferred to flash or fractionate thediluent transferred to the second (HMW) reactor hydrocyclone in order tominimize the hydrogen transferred to the downstream reactor.

However, where a plant has been designed to operate a single catalyst orproduct type, the volume and dimensions of each reactor may be optimisedindividually for the principal grades to be produced, and thus the tworeactors may be of different volumes and dimensions. These differentdimensions can be utilised in order to obtain the desired balance ofaverage activity between the two reactors in accordance with theinvention, thus providing greater freedom to vary other reactionparameters. Thus in order to achieve the desired ratio of averageactivity between the first (LMW) and second (HMW) reactors, in analternative embodiment of the LMW-HMW aspect of the invention, the first(LMW) reactor may have a space time yield (defined as production ofpolymer in kg/h per unit volume of reactor) greater than 150 kg/m³/h,more preferably greater than 200 kg/m³/h, and most preferably greaterthan 250 kg/m³/h. In this case the ratio of space time yield in thefirst (LMW) reactor to the second (HMW) reactor may be greater than 1.2,for example greater than 1.5. This may be achieved by designing thefirst (LMW) reactor with a volume that is no more than 90%, preferablybetween 30-70%, and more preferably approximately 40-60%, of the volumeof the second (HMW) reactor. In this case it is preferred that the ratioof length to diameter (L/D) of the first reactor is greater than 400,preferably between 400 and 800; additionally or alternatively, the ratioof L/D of the first reactor to L/D of the second reactor is greater than1.5, most preferably greater than 2.

The following comments apply to all aspects and embodiments of theinvention. One general benefit of the invention is that the optimisationof reactor average activity balance, space time yields and coolingrequirements, whilst at the same time minimising the C_(lights)concentration in the flash tank so as to avoid the need to recompressleads to improved efficiency. This invention can enable monomerefficiencies of less than 1.015, generally less than 1.01 and preferablyless then 1.006 to be achieved even when employing a space time yield ofat least 100 kg/m³/h, more preferably at least 150 kg/m³/h, mostpreferably at least 200 kg/m³/h in each reactor. By “monomer efficiency”is meant the weight ratio of ethylene+comonomer consumed to polymerproduced.

It is preferred in all aspects of the invention that the reactor is“liquid full”: in other words, there is substantially no gas- orvapour-containing headspace in the reactor.

A preferred type of reactor utilised for the polymerisations to whichall aspects of the invention are applicable is a loop reactor, which isa continuous tubular construction comprising at least two, for examplefour, vertical sections and at least two, for example four horizontalsections. The heat of polymerisation is typically removed using indirectexchange with a cooling medium, preferably water, in jackets surroundingat least part of the tubular loop reactor. The volume of one loopreactor in a multiple reactor system can vary but is typically in therange 10-200 m³. It is preferred that the polymerisation reactorutilised in the present invention is loop reactors, and further that theloop reactors are liquid full.

Typical pressures employed in the loop reactor are between 0.1-10 MPa g,preferably between 3 to 5 MPa g.

The process according to the invention applies to the preparation ofcompositions containing ethylene homopolymers and copolymers. Ethylenecopolymers typically comprise one or more alpha-olefins in a variableamount which can reach 12% by weight, preferably from 0.5 to 6% byweight, for example approximately 1% by weight.

The alpha mono-olefin monomers generally employed in such reactions areone or more 1-olefins having up to 8 carbon atoms per molecule and nobranching nearer the double bond than the 4-position. Typical examplesinclude ethylene, propylene, butene-1, pentene-1, hexene-1 and octene-1,and mixtures such as ethylene and butene-1 or ethylene and hexene-1.Butene-1, pentene-1 and hexene-1 are particularly preferred comonomersfor ethylene copolymerisation.

In one embodiment of the invention, the polymer is a polyethylene resinhaving a density of greater than 940 kg/m³ and an HLMI of from 1 to 100g/10 min, and comprising from 35 to 60 wt % of a first polyethylenefraction of high molecular weight and from 40 to 65 wt % of a secondpolyethylene fraction of low molecular weight, the first polyethylenefraction comprising a linear low density polyethylene having a densityof up to 935 kg/m³ and an HLMI of less than 1 g/10 min, and the secondpolyethylene fraction comprising a high density polyethylene having adensity of at least 960 kg/m³, preferably at least 965 kg/m³, and an MI₂of greater than 100 g/10 min, and the polyethylene resin.

Typical diluents for the suspensions in each reactor includehydrocarbons having 2 to 12, preferably 3 to 8, carbon atoms permolecule, for example linear alkanes such as propane, n-butane, n-hexaneand n-heptane, or branched alkanes such as isobutane, isopentane,isooctane and 2,2-dimethylpropane, or cycloalkanes such as cyclopentaneand cyclohexane or their mixtures. In the case of ethylenepolymerization, the diluent is generally inert with respect to thecatalyst, cocatalyst and polymer produced (such as liquid aliphatic,cycloaliphatic and aromatic hydrocarbons), at a temperature such that atleast 50% (preferably at least 70%) of the polymer formed is insolubletherein. Isobutane is particularly preferred as the diluent.

The operating conditions can also be such that the monomers act as thediluent as is the case in so called bulk polymerisation processes. Theslurry concentration limits in volume percent have been found to be ableto be applied independently of molecular weight of the diluent andwhether the diluent is inert or reactive, liquid or supercritical.Propylene monomer is particularly preferred as the diluent for propylenepolymerisation

Methods of molecular weight regulation are known in the art. When usingZiegler-Natta, metallocene and tridentate late transition metal typecatalysts, hydrogen is preferably used, a higher hydrogen pressureresulting in a lower average molecular weight. When using chromium typecatalysts, polymerization temperature is preferably used to regulatemolecular weight.

In commercial plants, the particulate polymer is separated from thediluent in a manner such that the diluent is not exposed tocontamination so as to permit recycle of the diluent to thepolymerization zone with minimal if any purification. Separating theparticulate polymer produced by the process of the present inventionfrom the diluent typically can be by any method known in the art forexample it can involve either (i) the use of discontinuous verticalsettling legs such that the flow of slurry across the opening thereofprovides a zone where the polymer particles can settle to some extentfrom the diluent or (ii) continuous product withdrawal via a single ormultiple withdrawal ports, the location of which can be anywhere on theloop reactor but is preferably adjacent to the downstream end of ahorizontal section of the loop. The operation of large diameter reactorswith high solids concentrations in the slurry minimises the quantity ofthe principal diluent withdrawn from the polymerisation loop. Use ofconcentrating devices on the withdrawn polymer slurry, preferablyhydrocylones (single or in the case of multiple hydrocyclones inparallel or series), further enhances the recovery of diluent in anenergy efficient manner since significant pressure reduction andvaporisation of recovered diluent is avoided. Increasing theconcentration of easily condensable components, for example throughaddition of fresh or recycle diluent, upstream of the hydrocyclone is afurther means of enhancing the operating window of the final reactor andreducing the concentration of monomer depressurised to the mediumpressure flash tank.

Where the final reactor of the multiple reactor system is a loopreactor, the withdrawn, and preferably concentrated, polymer slurry isdepressurised, and optionally heated, prior to introduction into aprimary flash vessel. The stream is preferably heated afterdepressurisation. As a consequence of the invention, the diluent and anymonomer vapours recovered in the primary flash vessel can be condensedwithout recompression. They are typically then recycled to thepolymerization process. Typically the pressure in the primary flashvessel is 0.5-2.5 MPa g, preferably 0.5-1.5 MPa g. The solids recoveredfrom the primary flash vessel are usually passed to a secondary flashvessel to remove residual volatiles.

The process according to the invention is relevant to all olefinpolymerisation catalyst systems, particularly those chosen from theZiegler-type catalysts, in particular those derived from titanium,zirconium or vanadium and from thermally activated silica or inorganicsupported chromium oxide catalysts and from metallocene-type catalysts,metallocene being a cyclopentadienyl derivative of a transition metal,in particular of titanium or zirconium.

Non-limiting examples of Ziegler-type catalysts are the compoundscomprising a transition metal chosen from groups IIIB, IVB, VB or VIB ofthe periodic table, magnesium and a halogen obtained by mixing amagnesium compound with a compound of the transition metal and ahalogenated compound. The halogen can optionally form an integral partof the magnesium compound or of the transition metal compound.

Metallocene-type catalysts may be metallocenes activated by either analumoxane or by an ionising agent as described, for example, in EP500944A (Mitsui Toatsu Chemicals).

Ziegler-type catalysts are most preferred. Among these, particularexamples include at least one transition metal chosen from groups IIIB,IVB, VB and VIB, magnesium and at least one halogen. Good results areobtained with those comprising:

from 10 to 30% by weight of transition metal, preferably from 15 to 20%by weight,

from 20 to 60% by weight of halogen, preferably from 30 to 50% by weight

from 0.5 to 20% by weight of magnesium, usually from 1 to 10% by weight,

from 0.1 to 10% by weight of aluminium, generally from 0.5 to 5% byweight, the balance generally consists of elements arising from theproducts used for their manufacture, such as carbon, hydrogen andoxygen. The transition metal and the halogen are preferably titanium andchlorine. Most preferred catalysts have the following composition:

Transition metal from 8 to 20% by weight

Magnesium content from 3 to 15% by weight

Chlorine content from 40 to 70% by weight

Aluminum content less than 5% by weight

Residual organic content less than 40% by weight

Polymerisations, particularly Ziegler catalysed ones, are typicallycarried out in the presence of a cocatalyst. It is possible to use anycocatalyst known in the art, especially compounds comprising at leastone aluminium-carbon chemical bond, such as optionally halogenatedorganoaluminium compounds, which can comprise oxygen or an element fromgroup I of the periodic table, and aluminoxanes. Particular exampleswould be organoaluminium compounds, of trialkylaluminiums such astriethylaluminium, trialkenylaluminiums such as triisopropenylaluminium,aluminium mono- and dialkoxides such as diethylaluminium ethoxide, mono-and dihalogenated alkylaluminiums such as diethylaluminium chloride,alkylaluminium mono- and dihydrides such as dibutylaluminium hydride andorganoaluminium compounds comprising lithium such as LiAl(C₂H₅)₄.Organoaluminium compounds, especially those which are not halogenated,are well suited. Triethylaluminium and triisobutylaluminium areespecially advantageous.

The chromium-based catalyst is preferred to comprise a supportedchromium oxide catalyst having a titania-containing support, for examplea composite silica and titania support. A particularly preferredchromium-based catalyst may comprise from 0.5 to 5 wt % chromium,preferably around 1 wt % chromium, such as 0.9 wt % chromium based onthe weight of the chromium-containing catalyst. The support comprises atleast 2 wt % titanium, preferably around 2 to 3 wt % titanium, morepreferably around 2.3 wt % titanium based on the weight of the chromiumcontaining catalyst. The chromium-based catalyst may have a specificsurface area of from 200 to 700 m²/g, preferably from 400 to 550 m²/gand a volume porosity of greater than 2 cc/g preferably from 2 to 3cc/g. Chromium-based catalysts may be used in conjunction withactivators such organometallic compounds of aluminium or of boron.Preferred are organoboron compounds such as trialkylborons in which thealkyl chains comprise up to 20 carbon atoms. Triethylboron isparticularly preferred.

If the catalyst employed is a metallocene catalyst, it preferablycomprises a bis-tetrahydroindenyl (THI) compound. Preferably thecatalyst system comprises (a) a metallocene catalyst componentcomprising a bis-tetrahydroindenyl compound of the general formula(IndH₄)₂R″MQ₂ in which each IndH₄ is the same or different and istetrahydroindenyl or substituted tetrahydroindenyl, R″ is a bridge whichcomprises a C₁-C₄ alkylene radical, a dialkyl germanium or silicon orsiloxane, or an alkyl phosphine or amine radical, which bridge issubstituted or unsubstituted, M is a Group IV metal or vanadium and eachQ is hydrocarbyl having 1 to 20 carbon atoms or halogen; and (b) acocatalyst which activates the catalyst component. Eachbis-tetrahydroindenyl compound may be substituted in the same way ordifferently from one another at one or more positions in thecyclopentadienyl ring, the cyclohexenyl ring and the ethylene bridge.Each substituent group may be independently chosen from those of formulaXR_(v) in which X is chosen from group IVB, oxygen and nitrogen and eachR is the same or different and chosen from hydrogen or hydrocarbyl offrom 1 to 20 carbon atoms and v+1 is the valence of X. X is preferablyC. If the cyclopentadienyl ring is substituted, its substituent groupsmust not be so bulky as to affect coordination of the olefin monomer tothe metal M. Substituents on the cyclopentadienyl ring preferably have Ras hydrogen or CH₃. More preferably, at least one and most preferablyboth cyclopentadienyl rings are unsubstituted. In a particularlypreferred embodiment, both indenyls are unsubstituted. R″ is preferablyan ethylene bridge which is substituted or unsubstituted. The metal M ispreferably zirconium, hafnium or titanium, most preferably zirconium.Each Q is the same or different and may be a hydrocarbyl or hydrocarboxyradical having 1-20 carbon atoms or a halogen. Suitable hydrocarbylsinclude aryl, alkyl, alkenyl, alkylaryl or aryl alkyl. Each Q ispreferably halogen. Ethylene bis(4,5,6,7-tetrahydro-1-indenyl) zirconiumdichloride is a particularly preferred bis tetrahydroindenyl compound.

Silica supported chromium catalysts are typically subjected to aninitial activation step in air at an elevated activation temperature.The activation temperature preferably ranges from 500 to 850° C., morepreferably 600 to 750° C.

In the process of the invention, the first reactor of the series issupplied with catalyst and the cocatalyst in addition to the diluent andmonomer, and each subsequent reactor is supplied with, at least,monomer, in particular ethylene and with the slurry arising from apreceding reactor of the series, this mixture comprising the catalyst,the cocatalyst and a mixture of the polymers produced in a precedingreactor of the series. It is optionally possible to supply a secondreactor and/or, if appropriate, at least one of the following reactorswith fresh catalyst and/or cocatalyst a first reactor.

1. Polymerisation process in which polyethylene is produced in slurry ina polymerisation reactor in the presence of a Ziegler Natta catalyst andan activator, and a stream or slurry containing the polymer is withdrawnfrom the reactor and transferred to a flash tank operating at a pressureand temperature such that at least 50 mol % of the liquid or non-polymercomponent of the stream entering the flash tank or slurry is withdrawnfrom the flash tank as a vapour and at least 98 mol % of the vapourwithdrawn from the flash tank is capable of being condensed at atemperature of between 15 and 50° C., without compression, wherein aby-product suppressor, which reduces the amount of by-product formed perunit of polyethylene produced by at least 10%, compared with anidentical polymerisation process where the by-product suppressor is notpresent, is used in the reactor, the molar ratio of the by-productsuppressor added to the reactor to titanium added to the reactor beingbetween 0.2 and
 1. 2. Process according to claim 1, wherein theby-product suppressor reduces the concentration in the reactor of atleast one of the components in the slurry which has a molecular weightbelow 50 by at least 5% compared with an identical process where theby-product suppressor is not present.
 3. Process according to claim 1,wherein the by-product suppressor is a halogenated hydrocarbon. 4.Polymerisation process in which polyethylene is produced in slurry in apolymerisation reactor in the presence of a Ziegler Natta catalyst andan activator, and a stream or slurry containing the polymer is withdrawnfrom the reactor and transferred to a flash tank operating at a pressureand temperature such that at least 50 mol % of the liquid or non-polymercomponent of the stream entering the flash tank or slurry is withdrawnfrom the flash tank as a vapour and at least 98 mol % of the vapourwithdrawn from the flash tank is capable of being condensed at atemperature of between 15 and 50° C., without compression, wherein ahalogenated hydrocarbon is present in the slurry, the molar ratio of theby-product suppressor added to the reactor to titanium added to thereactor being between 0.2 and
 1. 5. Process according claim 1 or 4,wherein a stream containing the polymer is withdrawn from the reactorand transferred to a flash tank operating at a pressure and temperaturesuch that at least 50 mol % of the non-polymer component of the streamentering the flash tank is withdrawn from the flash tank as a vapour. 6.Process according to claim 1 or 4, wherein the concentration in thestream entering the flash tank of components having a molecular weightbelow 50 g/mol, C_(lights) (mol %), satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) where T_(c)and P_(c) are respectively the temperature (in ° C.) and pressure (MPag) at the location where the vapour withdrawn from the flash tank iscondensed, and C_(H2) and C_(Et) are the molar concentrations in theflash tank of hydrogen and ethylene respectively.
 7. Process accordingto claim 6, wherein the concentration in the reactor of componentshaving a molecular weight below 50 also satisfies the equation_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) whereC_(lights), C_(H2), and C_(Et) in this case are the concentrations ofcomponents having a molecular weight below 50, hydrogen and ethylenerespectively in the reactor.
 8. Process according to claim 6, whereinthe concentration in the slurry entering the flash tank of hydrogen,ethylene and components having a molecular weight below 50 is the sameas the concentration of hydrogen, ethylene and components having amolecular weight below 50 in the second reactor.
 9. Process according toclaim 1 or 4, wherein at least 80 mol % of the liquid component of theslurry is withdrawn from the flash tank as a vapour.
 10. Processaccording to claim 1 or 4, wherein the polyethylene is a bimodal polymermade in two reactors in series comprising a first reactor and a secondreactor, and the by-product suppressor is added to at least one of thereactors.
 11. Process according to claim 10, wherein a high molecularweight (HMW) polymer is made in suspension in the first reactor and alow molecular weight (LMW) polymer is made in suspension in the secondreactor in the presence of the first polymer.
 12. Process according toclaim 10, wherein a high molecular weight (HMW) polymer is made insuspension in the first reactor and a low molecular weight (LMW) polymeris made in suspension in the second reactor in the presence of the firstpolymer, and the ratio of the average activity in the second LMW reactorto the average activity in the first HMW reactor is from 0.25 and 1.5,where average activity in each reactor is defined as the rate ofpolyethylene produced in the reactor (kgPE/hr)/[ethylene concentrationin the reactor (mol %)×residence time in the reactor (hours)×feed rateof catalyst into the reactor (g/hr)].
 13. Process according to claim 10,wherein a high molecular weight (HMW) polymer is made in suspension inthe first reactor and a low molecular weight (LMW) polymer is made insuspension in the second reactor in the presence of the first polymer,and the ratio of ethylene concentration (in mol %) in the second reactorto that in the first reactor is 5 or less.
 14. Process according toclaim 10, wherein a high molecular weight (HMW) polymer is made insuspension in the first reactor and a low molecular weight (LMW) polymeris made in suspension in the second reactor in the presence of the firstpolymer, and the concentration of ethylene in the second reactor is lessthan 8 mol %.
 15. Process according to claim 1 or 4, wherein thepolyethylene is a multimodal polyethylene having a shear ratio of atleast 15, where “shear ratio” is the ratio of the high load melt indexHLMI of the polyethylene to the MI₅ of the polyethylene, both beingmeasured according to ISO Standard 1133 at a temperature of 190° C. 16.Process according to claim 1, wherein the by-product suppressor reducesthe absolute amount of by-product formed per unit of polyethyleneproduced by at least 10% compared with an identical polymerisationprocess where the by-product suppressor is not present.
 17. Processaccording to claim 3, wherein the by-product suppressor is achloromethane of the formula CH_(x)Cl_(4-x) where x is an integer from 1to
 3. 18. Process according to claim 17, wherein the by-productsuppressor is chloroform, CHCl₃.
 19. Process according to claim 1 or 4,wherein at least 98 mol % of the vapour withdrawn from the flash tank iscapable of being condensed at a temperature of between 15 and 40° C. 20.Process according to claim 4, wherein the halogenated hydrocarbon is achloromethane of the formula CH_(x)Cl_(4-x) where x is an integer from 1to 3.